Hydrocracking Processes

Distillate hydrocracking is a refining process for conversion of heavygas oils and heavy diesels or similar boiling-range heavy distillates intolight distillates (naphtha, kerosene, diesel, etc.) or base stocks for lubri-cating oil manufacture. The process consists of causing feed to react withhydrogen in the presence of a catalyst under specified operating condi-tions: temperature, pressure, and space velocity.

HYDROCRACKING REACTIONS

DESULFURIZATION

The feedstock is desulfurized by the hydrogenation of the sulfur con-taining compounds to form hydrocarbon and hydrogen sulfide. The H2S isremoved from the reactor effluent leaving only the hydrocarbon product.The heat of reaction for desulfurization is about 60 Btu/scf of hydrogenconsumed:

CATALYST

Thiophene Paraffin Hydrogen Sulphide

DENITRIFICATION

Nitrogen is removed from feedstock by the hydrogenation of nitrogen-containing compounds to form ammonia and hydrocarbons. Ammonia islater removed from the reactor effluent, leaving only the hydrocarbons inthe product. The heat of reaction of the denitrification reactions is about

67-75 Btu/scf of hydrogen consumed, but the amount of nitrogen in thefeed is generally very small, on the order of a few parts per million, andhence its contribution to overall heat of reaction is negligible:

Amine Paraffin Ammonia

OLEFIN HYDROGENATION

The hydrogenation of olefins is one of the most rapid reaction takingplace, and therefore almost all olefins are saturated. The heat of reactionis about 140 Btu/scf of hydrogen consumed. Olefin content is generallysmall for straight-run products, but for stocks derived from secondary/thermal processes such as coking, visbreaking, or resid hydrocracking(H-OIL* etc.), it can contribute a considerable amount of heat liberated inthe hydrocracker reactor.

RCH2CH = CH2 + H2 = RCH2CH2CH3

Olefin Paraffin

SATURATION OF AROMATICS

Some of the aromatics in the feed are saturated, forming naphthenes.Saturation of aromatics accounts for a significant proportion of both thehydrogen consumption and the total heat of reaction. The heat of reactionvaries from 40 to 80 Btu/scf of hydrogen consumed, depending on thetype of aromatics being saturated. In general, higher reactor pressure andlower temperature result in a greater degree of aromatic saturation:

AROMATIC NAPHTHENE

*H-0IL is a commercial processed for resid hydrocracking/resid desulfurisation, licensed byHydrocarbon Research Inc. (USA).

HYDROCRACKING OF LARGE MOLECULES

Hydrocracking of large hydrocarbon molecules into smaller moleculesoccurs in nearly all processes carried out in the presence of excesshydrogen. These reactions liberate about 50Btu/scf of hydrogen con-sumed. The heat released from the hydrocracking reactions contributesappreciably to the total heat liberated in the reactor. Cracking reactionsinvolving heavy molecules contribute to lowering the specific gravity andforming light products, such as gas and light naphtha, in the hydrocrackerproducts.

An example of a hydrocracking reaction is

RCH2CH2CH2CH3 + H 2 - RCH3 + CH3CH2CH3

The yield of light hydrocarbons is temperature dependent. Therefore,the amount of light end products produced increases significantly as thereactor temperature is increased to compensate for a decrease in catalystactivity toward the end of run conditions.

FEED SPECIFICATIONS

Hydrocracker feed is typically heavy diesel boiling above the saleablediesel range or vacuum gas oil stream originating from the crude andvacuum distillation unit, atmospheric resid desulfurizers, coker units,solvent deasphalting units, and the like. The hydrocracking catalyst isvery sensitive to certain impurities, such as nitrogen and metals, and thefeed must conform to the specifications laid down by the catalyst man-ufacturers to obtain a reasonable catalyst life.

FEED NITROGEN

Nitrogen in the feed neutralizes catalyst acidity. Higher nitrogen in thefeed requires slightly more severe operating conditions, particularly thetemperature, and causes more rapid catalyst deactivation.

FEED BOILING RANGE

A higher-than-designed feed distillation end point accelerates catalystdeactivation and requires higher reactor temperatures, thus decreasingcatalyst life.

The feed properties have little direct effect on light product yield, butthey affect the catalyst temperature required to achieve the desired con-version. The yield of light gases (C4-) and naphtha boiling-range materialis increased when the catalyst temperature is increased.

ASPHALTENES

In high cut point vacuum distillation, there is always a possibility thatexcessive high molecular weight, multiring aromatics (asphaltenes) canbe found in vacuum gas oil distillates. In addition to causing excessivecatalyst poisoning, asphaltenes may be chemically combined with thecatalyst to deactivate the catalyst permanently.

METALS

Metals, particularly arsenic, and alkalies and alkaline earth deposit inthe catalyst pores reduce catalyst activity. Common substances that cancarry metallic catalyst contaminants include compounded lubricating oilsor greases, welding fluxes, and gasketing.

Iron carried in with the feed is likely to be the most troublesomemetallic catalyst contaminant. It may be chemically combined with heavyhydrocarbon molecule, or it may exist as suspended particulate matter. Ineither case, it not only deactivates the catalyst but also plugs the catalystinterstices such that excessive pressure drop develops. Normally, thisplugging appears as a crust at the top of the first catalyst bed.

CHLORIDES

The feed may contain trace amounts of organic and inorganic chlor-ides, which combine with ammonia produced as a result of denitrificationreactions to form very corrosive deposits in the reactor effluent exchangerand lines.

OXYGEN

Oxygenated compounds, if present in the feed, can increase deactiva-tion of the catalyst. Also, oxygen can increase the fouling rate of the feedeffluent heat exchangers.

CATALYST

Hydrocracking reactions can be divided into two groups: (1) desulfur-ization and denitrification—hydrogenation of polyaromatics and mono-aromatics—are favored by the hydrogenating function of the catalyst(metals) and (2) hydrodealkylation, hydrodecyclization, hydrocracking,and hydroisomerization reactions are promoted by the acidic function ofthe catalyst (support). The support function is affected by the nitrogencontent of the feed.

The catalyst employed in hydrocracking is generally of the type (Ni-Co-Fe), (Mo-W-U) on a silica/alumina support. The ratio of alumina tosilica is used to control the degree of hydrocracking, hydrodealkylation,hydroisomerization, and hydrodecyclization. Cracking reactions increasewith increasing silica content of the catalyst. Metals, in the form ofsulfide, control the desulfurization, denitrification, and hydrogenation ofolefins, aromatics, and the like.

The choice of catalyst system depends on the feedstock to be treatedand the products required. Most of the time, the suitable system isobtained by the use of two or more catalysts with different acidic andhydrogenation functions. The reactor may also contain a small amount, upto 10%, of desulfurization and denitrification catalyst in the last bed ofthe reactor.

PROCESS CONFIGURATION

Hydrocracker units can be operated in the following possible modes:single-stage (once-through-mode) operation, single-stage operation withpartial or total recycling, and two-stage operation. These operation modesare shown in Figures 3-1 and 3-2.

The choice of the process configuration is tied to the catalyst system.The main parameters to be considered are feedstock quality, the productslate and qualities required, and the investment and operating costs ofthe unit.

SINGLE-STAGE OPERATION

This operating mode has large effect on the product yield and quality.Single-stage operation produces about 0.3 bbl naphtha for every barrel ofmiddle distillate. The single stage scheme is adapted for conversion of

FRESH FEED REACTORSECTION DISTILLATION

SECTION

C4

NAPHTHA

KEROSENE

DIESEL

FRACTIONATOR BTMS.

ONCE-THROUGH MODE

FRESH FEED REACTORSECTION

DISTILLATIONSECTION

C4

NAPHTHA

KEROSENE

DIESEL

DESULFURISEDGAS OILFRACTIONATOR BOTTOMS

PARTIAL-RECYCLE MODE

Figure 3-1. Hydrocracker operation, once-through and partial-recycle modes.

FRESH FEED

R E A C T O RSECTION

DISTILLATIONSECTION

C 4

NAPHTHA

KEROSENE

DIESEL

FRACTIONATOR BOTTOMS

FIRST STAGE

R E A C T O RSECTION

2ND STAGE

DISTILLATIONSECTJON

2ND STAGE

C 4

NAPHTHA

KEROSENE

DIESEL

RECYCLE STREAM

{OPTIONAL J

FRACTIONATOR BOTTOMS

Figure 3-2. Two-stage hydrocracking process.

vacuum gas oils into middle distillate and allows for high selectivity. Theconversion is typically around 50-60%. The unconverted material is lowin sulfur, nitrogen, and other impurities and is used as either feed for fluidcatalytic cracking units (FCCU) or a fuel oil blending component.

The single-stage process may be operated with partial or total recyc-ling of the unconverted material. In total recycling, the yield of naphtha isapproximately 0.45 bbl per barrel of middle distillate products. In thesecases, the fresh feed capacity of the unit is reduced. Thus, increasedconversion is achieved basically at the cost of unit's fresh feed capacityand a marginally increased utility cost. The partial recycling mode ispreferable to total recycling to extinction, as the latter results in thebuildup of highly refractory material in the feed to the unit, resulting inhigher catalyst fouling rates.

TWO-STAGE OPERATION

In the two-stage scheme, the unconverted material from the first stagebecomes feed to a second hydrocracker unit. In this case, the feed is alreadypurified by the removal of sulfur, nitrogen, and other impurities; and thesecond-stage can convert a larger percentage of feed with better productquality.

A heavy gas oil feed contains some very high-boiling aromatic mol-ecules. These are difficult to crack and, in feed recycling operation, tendto concentrate in the recycle itself. High concentration of these moleculesincrease the catalyst fouling rate. In a two-stage operation, the first stage isa once-through operation; hence, the aromatic molecules get no chance toconcentrate, since there is no recycle. The first stage also reduces theconcentration of these molecules in the feed to the second stage; there-fore, the second stage also sees lower concentration of these high-boilingaromatic molecules.

The two-stage operation produces less light gases and consumes lesshydrogen per barrel of feed. Generally, the best product qualities (lowestmercaptans, highest smoke point, and lowest pour point) are producedfrom the second stage of the two-stage process. The poorest qualities arefrom the first stage. The combined product from the two stages is similarto that from a single stage with recycling for the same feed quality.

The two-stage scheme allows more flexible adjustment of operating con-ditions, and the distribution between the naphtha and middle distillate is moreflexible. Compared to partial and total recycling schemes, the two-stagescheme requires a higher investment but is overall more economical.

PROCESS FLOW SCHEME

The oil feed to the reactor section consists of two or more streams (seeFigure 3-3). One stream is a vacuum gas oil (VGO) feed from the storagetank and the other stream may be the VGO direct from the vacuumdistillation unit. Also, there may be an optional recycling stream consist-ing of unconverted material from fractionator bottom. The combined feedis filtered in filters F-Ol to remove most of the particulate matter thatcould plug the catalyst beds and cause pressure drop problems in thereactor. After the oil has passed through surge drum V-02, it is pumped tothe reactor system pressure by feed pump P-Ol.

Hydrogen-rich recycled gas from the recycling compressor is combined withoil feed upstream of effluent/feed exchangers E-01/02. The oil gas stream thanflows through the tube side of exchanger 02A and 02B, where it is heated byexchange with hot reactor effluent. Downstream of the feed effluent exchan-gers, the mixture is further heated in parallel passes through reactor feed heaterH-Ol. The reactor inlet temperature is controlled by the Temperature Recorderand Controller (TRC) by controlling the burner fuel flow to the furnace.

A portion of the oil feed is by passed around the feed effluent exchan-ger. This bypass reduces the exchanger duty while maintaining the dutyof reactor feed heater H-Ol at a level high enough for good control ofreactor inlet temperature. For good control, a minimum of 50-750Ftemperature rise across the heater is required.

Makeup hydrogen is heated on the tube side of exchanger E-Ol by thereactor effluent. This makeup hydrogen then flows to the reactor.

Hydrocracker reactor V-Ol is generally a bottle-type reactor. Themakeup hydrogen after preheating in exchangers E-Ol flows up throughthe reactor in the annular space between the reactor outside shell and aninside bottle. The hydrogen acts as a purge to prevent H2S from accumu-lating in the annular space between the bottle and outside shell. It alsoinsulates the reactor shell.

After the makeup hydrogen has passed upward through the reactor, itcombines with the recycled gas and the heated oil feed from the feedheater in the top head of the reactor. The hot, vaporized reaction mixturethen passes down the reactor. Cold quenching gas from the recyclingcompressor is injected to the reactor between the catalyst beds to limit thetemperature rise produced by exothermic reactions.

The reactor is divided among a number of unequal catalyst beds. Thisis done to give approximately the same temperature rise in each catalyst

BLEED

TO RELIEF

i LEAN DEA FROMREGENERATOR

RICH DEA TOREGENERATOR

OFF-GASTOHPAMINE CONTACTOR

TO H2S STRIPPERPREFLASH(DISTILLATION SECTION)

SOURWATERTO TREATMENT

H2S STRIPPER FEEDFROM V-OMO(DISTILLATION SECTION)

TO H2S STRIPPER

V-01-04HS ABSORBER

OIL SKIM

V-01-03K.O. DRUM

V-01-06LOW PRESSURESEPERATOR

550 PS G

E-01-03REACTOR EFFLUENT/H2S STRIPPED FEEDEXCHANGER

V-01-08INJECTION!WATERSURGE IDRUM

B.F.W

N2 POLYSULFIDE

V-OI-01POLYSULFIDEDRUM

STEAMSTEAM

C-01-01RECYCLECOPPRESSOR

V-01-05K.O. DRUM

TO DEA FLASHDRUM

V-01-02HIGH PRESSURESEPARATOR

2d50 PS G

E-01-04REACTOR EFFLUENTAIR COOLER

NITROGENPURGE

2715 PSIG185°F

2515 PSIG800°F58% VAPOR1

E-01-02FEED EFFLUENTEXCHANGER

E-01-01MAKEUP HYDROGENREACTOR EFFLUENTEXCHANGER

V-01-01REACTOR

2670 PSIG717°F

800 PSIG2200F

MAKEUPHYDROGEN

H-01-01FEED HEATER

F-01-01FEEDFILTERS

INERT GASBLANKET

V-01-07FEED SURGE DRUM(15 MINUTES)

P-01-01FEED PUMP

FEED FROM STORAGE

FEED FROM UNITS

BOTTOM RECYCLEFROM FRACTIONATOR

Figure 3-3. Distillate hydrocracker (reactor section). K.O. = knockout.

bed and limit the temperature rise to 500F. Thus, the first and second bedsmay contain 10 and 15% of the total catalyst, while the third and fourthbeds contain 30 and 45% of the total catalyst.

Reactor internals are provided between the catalyst beds to ensurethorough mixing of the reactants with quench and ensure gooddistribution of vapor and liquid flowing to each bed. Good distribu-tion of reactants is of utmost importance to prevent hot spots andmaximize catalyst life.

Directly under the reactor inlet nozzle is a feed distributor cone insidea screened inlet basket. These internal elements initiate feed streamdistribution and catch debris entering the reactor. Below the inlet basket,the feed stream passes through a perforated plate and distributor tray forfurther distribution before entering the first catalyst bed.

Interbed internal equipment consists of the following:

• A catalyst support grid, which supports the catalyst in the first bed,covered with a wire screen.

• A quench ring, which disperses quenching gas into hot reactantsfrom the bed above.

• A perforated plate for gross distribution of quenched reaction mix-ture.

• A distribution tray for final distribution of quenched reaction mix-ture before it enters the next catalyst bed.

• A catalyst drain pipe, which passes through interbed elements andconnects each catalyst bed with the one below it.

To unload the catalyst charge, the catalyst from the bottom bed isdrained through a catalyst drain nozzle, provided in the bottom head ofthe reactor. Each bed then drains into the next lower bed through the beddrain pipe, so that nearly all the catalyst charge can be removed witha minimum of effort.

Differential pressure indicators are provided to continuously measurepressure drop across top reactor beds and the entire reactor.

The reactor is provided with thermocouples located to allow obser-vation of catalyst temperature both axially and circumferentially. Thermo-couples are located at the top and bottom of each bed. The temperaturemeasured at the same elevation but different circumferential positionsin the bed indicate the location and extent of channeling throughthe beds.

EFFLUENT COOLING

The reactor effluent is at 800-8500F, start of run (SOR) to end of run(EOR), at reactor outlet. The reactor effluent is cooled in makeup hydro-gen/reactor effluent exchanger E-Ol and reactor effluent/H2S stripperfeed exchanger E-03. The reactor effluent is further cooled in reactoreffluent/air cooler E-04 to about 1400F.

WATER AND POLYSULFIDE INJECTION

Condensate is injected into the reactor effluent just upstream of efflu-ent air cooler. The function of the injection water is to remove ammoniaand some H2S from the effluent. The effluent temperature at the injectionpoint is controlled to prevent total vaporization of the injected water andpreclude deposition of solid ammonium bisulfide.

Trace amounts of cyanide ion in the reactor effluent contribute tocorrosion in the effluent air cooler. A corrosion inhibitor such as sodiumpolysulfide is also injected to prevent cyanide corrosion.

HIGH-PRESSURE SEPARATOR

The high-pressure separator, V-02, temperature is controlled atapproximately 1400F by a temperature controller, which adjusts thepitch of half the fans of the air cooler. The temperature of the separatoris closely controlled to keep the downstream H2S absorber temperaturefrom fluctuating. The hydrogen purity is lower at higher temperatures.However, at lower temperature, poor oil/water separation occur in theseparator drum.

LOW-TEMPERATURE SEPARATOR

The high-temperature separator liquid is depressurized through a HPseparator level control valve to 550 psig and flashed again to low-pressureseparator V-06. The low-pressure separator overhead vapors flow to high-pressure amine contractor V-04. The hydrocarbon stream leaving theseparator is fed to H2S stripper V-Il via a stripper preflash drum. Thesour water drawn from the low-pressure separator is sent to sour watertreating facilities.

RECYCLE GAS ABSORBER

DEA (diethanolamine) absorption is used to remove H2S from therecycled gas in recycle gas absorber V-04. H2S gas is absorbed by theDEA solution because of the chemical reaction of DEA with H2S. Typicalproperties of DEA are shown in Table 8-19.

The amount of H2S that can react depends on the operating conditions.The low temperature, high pressure, and high H2S concentration in theH2S absorber favor the reaction. In the DEA regeneration facilities, hightemperature and low pressure are used to reverse the reaction and stripH2S from the DEA solution.

About 90% of the H2S formed by the desulfurization reactions isremoved from the recycled gas in a high-pressure absorber by scrubbingthe gas with aqueous diethanolamine solution. The absorber is a verticalvessel packed with stainless steel ballast rings. Recycled gas flowsthrough a support plate and upward through the packing. A lean DEAsolution from the DEA regenerator enters the top of the absorber throughan inlet distributor and flows downward through the packing. Rich DEAfrom the bottom of the absorber is sent to the H2S recovery unit.

RECYCLE GAS COMPRESSOR

Recycled gas is circulated by recycle gas compressor C-Ol, driven bya steam turbine. The largest portion of the recycled gas stream joins theoil feed stream upstream of feed effluent exchangers. A portion of the gasstream from the recycling compressor flows on temperature control tointerbed quenching.

DISTILLATION SECTION

The purpose of distillation section (see Figure 3-4) is to remove H2Sand light ends from the first-stage reactor effluent and fractionate theremaining effluent into naphtha, kerosene, and diesel cuts. The bottomstream is either fed to the second stage of the hydrocracker, recycled toextinction with the fresh feed, or withdrawn as product.

Hydrocarbon liquid flows to H2S stripper V-Il from stripper preflashdrum V-IO. The preflash drum removes some of the light ends and H2Sfrom the low-pressure separator oil before it is stripped and fractionated.The stripper column contains packed sections below the feed plate andtwo sieve trays above the feed inlet. Stripping is achieved with steam,

TO AMINECONTRACTOR

VENTTORELIEF

FUEL GAS

CONDENSATE

NAPHTHA

FRACTIONATORBOTTOMS

KEROSENE

HEAVY DIESEL

TO RECYCLING/2ND

STAGE

V-01-14REFLUX DRUM

P-01-06

V-01-15SIDE CUT

STRIPPER

P-01-05

E-01-12

150 PSIG

STEAM

390°F

FROM H2S

E-01-15

TOFRACTIONATORCOLUMN

SIDE CUT

6 TRAYS

STEAM

P-01-09

TO S FROM

E-01-12SIDE STRIPPER

REBOILER

10 PSIG

155°F

E-01-08

440°F

565T

20 PSIG

695°F

675°F

P-01-10

E-01-15

190°F

TOOILYWATER DRAIN

CONDENSATE

FRACTIONATOR

32 TRAYS

P-01-07H-01-02

150 PSIG STEAM 600-F

515°F

E-01-18

490°F

10 PSIG

STM

V-01-12REFLUX DRUM

110°F

P-01-04

215°F105 PSIG

V-01-17WATER DRAWOFF

H2S STRIPPER

FROM FRACTIONATOR

E-Ot-OB

DIESEL PRODUCTE-01-16

330°F110 PSIG

STRIPPER

PREFLASHSEPERATOR LIQUID

198-F150 PSIG

137-F

WATERTOOILY WATER

REACTOR

350-F99% VAPORREACTOR SECTION

Figure 3-4. Distillate hydrocracker (distillation section).

which removes H2S and light ends. The stripper overhead vapor ispartially condensed in air cooler E-07 and a trim water cooler then flashedin reflux drum V-12. The sour gas from the reflux drum is sent to a low-pressure amine contactor. The condensed hydrocarbon liquid is refluxedback to the stripper. H2S stripper bottoms are sent to product fractionatorV-13 after heat exchange with diesel in E-16, circulating reflux in E-18,and fractionator bottoms in E-18. The fractionator feed is then brought tocolumn temperature by heating in the feed heater H-02. After heating, thepartially vaporized fractionator feed is introduced into the flash zone ofproduct fractionator V-12. In the flash zone, the vapor and liquid separate.The vapor passes up through the rectifying section, containing approxi-mately 27 trays.

Heat is removed from the fractionator column in the overhead con-denser and in a circulating reflux heat removal system. Vapor leaving thetop tray of the column is condensed in overhead condensers. The con-densed overhead vapor is separated into hydrocarbon and water phases.Part of the hydrocarbon is recovered as overhead product, and the rest issent back to the column as reflux to ensure good separation.

The portion of the column below the flash zone contains five trays.Superheated steam is injected below the bottom tray. As the steam passesup through the stripping section, it strips light components from theresidual liquid from the flash zone.

OPERATING CONDITIONS

The operating conditions for single-stage hydrocracking are shown inTable 3-1. The yields and qualities for once-through operation, two-stagehydrocracking, and one-stage operation with partial recycling of uncon-verted material are shown in Tables 3-2 to 3-5. It must be stressed that theyields depend on the catalyst composition and process configurationemployed, and these can vary significantly.

TEMPERATURE

The typical hydrocracker reactor operates between 775-825°F and2600 psig reactor inlet pressure. The high temperatures are necessaryfor the catalyst to hydrocrack the feed. The high reactor pressure is

Table 3-1Single-Stage Hydrocracker Operating Conditions

OPERATING PARAMETERS UNITS

CATALYST AVERAGE TEMPERATURE 0F 775SPACE VELOCITY, LHSV hr~l 1.72REACTOR INLET PRESSURE psig 2600REACTOR PRESSURE DROP psi 50HYDROGEN PARTIAL PRESSURE, INLET psi 2000HYDROGEN CHEMICAL CONSUMPTION scf/bbl 1150MAKEUP + RECYCLE AT REACTOR INLET, SOR scf/bbl 5000MAKEUP H2 PURITY VOL% 95HP SEPARATOR TEMPERATURE 0F 140HP SEPARATOR PRESSURE, SOR psig 2415BLEED RATE, SOR (100% H2) scf/bbl 200RECYCLE COMPRESSOR SUCTION PRESSURE psig 2390RECYCLE COMPRESSOR DISCHARGE PRESSURE psig 2715

necessary for catalyst life. Higher hydrogen partial pressure increasescatalyst life. To keep the hydrogen partial pressure at a high level, highreactor pressure and high hydrogen content of the reactor feed are neces-sary. To accomplish this, an excess of hydrogen gas is recycled throughthe reactor. A makeup hydrogen stream provides hydrogen to replace thehydrogen consumed chemically in hydrocracking, olefin, aromaticssaturation, and the hydrogen lost to atmosphere through purge or dis-solved in oil. A cold hydrogen-rich gas is injected between the catalystbeds in the reactor to limit the temperature rise caused by exothermichydrocracking reactions.

In the hydrocracking process, the feed rate, operating pressure, andrecycle gas rate are normally held constant. The reactor temperature isthe only remaining variable requiring close control to achieve therequired liquid feed conversion.

As the catalyst activity declines with time on stream due to catalystfouling, it becomes necessary to increase the reactor temperature tomaintain the original liquid feed conversion rate. This rate of increaseof reaction temperature with time is called the fouling rate. Additionaltemperature variation may be required to compensate for changes inreactor feed rate or feed properties, gas/oil ratio, hydrogen partial pres-sure, and the like (see Figure 3-5).

Table 3-2Hydrocracker Feed and Product Qualities

DIESELKEROSENEHEAVY

NAPHTHALIGHT

NAPHTHAFEEDPROPERTY

34176

D-86440/-

545/590610

640/685-/725

26492510

5

202560

9

38124

D-86350/-

375/410445

475/515-/550

205525

<10<5

15-50-80-55

47

2

53106

D-86215/-

225/245260

270/290-/325

305416

<10<5

61

80

D-86110/-

115/125135

150/170-/195

76204

76

22.6174

D-I160590/-

715/790840

885/940960/1010

162262

25000

970

75

59

SINGLE-STAGE OPERATIONAPIANILINE POINT 0FA.STM

IBP/510/305070/9095/FBP

COMPOSITION, LV%PARAFFINSNAPHTHENESAROMATICSSULFUR, ppmwMERCAPTANS, ppmNITROGEN, ppmwSMOKE POINT mmFREEZE POINT 0FPOUR POINT 0FCLOUD POINT 0FDIESEL INDEXOCTANE CLEARVISCOSITY, CST, 1000F

Table 3-2Continued

DIESELKEROSENEHEAVY

NAPHTHALIGHT

NAPHTHAFEEDPROPERTY

38193

D-86485/-

560/585600

620/655/705

4843

98

1.5

68

73.2

42.2148

D-86330/-

350/390420

460/520-/550

3845176

222

-50-65

56124

D-86195/-

215/230250

265/290-/320

4547

10

54

80

D-86110/-

115/125135

150/170-/195

7721

<5

80

20.9175

D-I160740/760775/835

875915/975

975/1035

31000

1050

90

SINGLE-STAGE WITHRECYCLE

APIANILINE POINT 0FASTM

IBP/510/305070/9095/FBP

COMPOSITION, LV%PARAFFINSNAPHTHENESAROMATICSSULFUR, ppmwMERCAPTANS, ppmNITROGEN, ppmwSMOKE POINT, mmCETANE NUMBERFREEZE POINT 0FPOUR POINT 0FCLOUD POINT 0FDIESEL INDEXOCTANE CLEAR

37190

460/520545/590

610635/675695/720

4242166

3

101570

8

40169

375/400415/475

540595/655675/700

394615

<5<5

2

-15-10

68

4

43147

345/365375/410

446475/515530/550

355015

<5<5<120

-54-70-65

63

2

56126

210/-225/240

255270/295305/325

4349

<5<5

54

81

110/-115/125

135150/170

-/195

8315

<5<5

75

TWO-STAGE OPERATIONAPIANILINE POINT 0FASTM D-86

IBP/510/305070/9095/FBP

COMPOSITION, LV%PARAFFINSNAPHTHENESAROMATICSSULFUR, ppmwMERCAPTANS, ppmNITROGEN, ppmwSMOKE POINT, mmFREEZE POINT 0FPOUR POINT 0FCLOUD POINT 0FDIESEL INDEXOCTANE CLEARVISCOSITY, CST, 1000F

Table 3-3Distillate Hydrocracker Yields

TWO-STAGEOVERALL

PARTIALRECYCLE

ONCETHROUGHSTREAM

1.00000.0345

0.0345

0.03910.01270.02210.21320.26360.4365

0.0473

1.0345

1.00000.0293

1.0293

0.02570.00690.02430.16370.18610.3955

0.2271

1.0293

1.00000.0225

1.0225

0.01660.00700.02160.09690.13720.29610.4471

1.0225

VGO FEEDHYDROGEN

TOTAL FEED

GASESHPGASACID GASCRACKED NAPHTHAKEROSENEDIESELHEAVY DIESEL (ONCE-THROUGH MODE)HEAVY DIESEL (PARTIAL-RECYCLE MODE)BLEED FROM SECOND STAGE

TOTAL PRODUCTS

NOTE: ALL FIGURES ON WAV BASIS. WAV = WEIGHTAVEIGHT BASIS.

Table 3-4Utility Consumption (per Ton Feed)

PARTIALUTILITY UNITS SINGLE STAGE RECYCLE MODE

FUEL GAS mmBtu 0.6 0.76POWER kWhr 18 23STEAM mmBtu 0.12 0.15DISTILLEDWATER MIG* 0.016 0.02COOLINGWATER MIG* 0.33 0.42

*MIG = 1000 IMPERIAL GALLONS.

Table 3-5Mild Hydrocracker Operating Conditions

OPERATING PARAMETERS UNITS

CATALYST AVERAGE TEMPERATURE 0F 775SPACE VELOCITY, LHSV hr"1 1.4REACTORINLETPRESSURE psig 1051REACTOR PRESSURE DROP psi 38HYDROGEN PARTIAL PRESSURE, INLET psi 749HYDROGEN CHEMICAL CONSUMPTION scf/bbl 358MAKEUP + RECYCLE AT REACTOR INLET, SOR scf/bbl 2766MAKEUP H2 PURITY VOL% 92HP SEPARATOR TEMPERATURE 0F 140HP SEPARATOR PRESSURE, SOR psig 850BLEED RATE, SOR (100% H2) scf/bbl 10.5RECYCLE COMPRESSOR SUCTION PRESSURE psig 820RECYCLE COMPRESSOR DISCHARGE PRESSURE psig 1070LEAN DEA TEMPERATURE 0F 150

CATALYST AVERAGE TEMPERATURE

The catalyst average temperature (CAT) is determined by the follow-ing equation:

RE

LA

TIV

E C

ATA

LYS

TF

OU

LIN

G R

ATE

DESIGN

GAS/OIL RATIO SCF/BBL

Figure 3-5. The effect of the gas/oil ratio on the catalyst fouling rate.

where T1 is the bed inlet temperature; T0 is the bed outlet temperature;and A1, A2, A3, and A4 are the volume fractions of the total reactorcatalyst in the individual bed. A typical hydrocracker reactor temperatureprofile is shown in Figure 3-6.

CATALYST FOULING RATE

The design of a hydrocracker unit is based on a specified conversionrate of the feed and a specified catalyst life, usually 2-3 years betweencatalyst regeneration. During the course of the run, the activity of thecatalyst declines due to coke and metal deposits, and to maintain thedesign conversion rate, the temperature of the catalyst has to be increased.The catalyst manufacturers specify a maximum temperature, called theend of run temperature, which signifies the EOR condition. When thistemperature is reached, the catalyst must be regenerated or discarded.

The rate of increase in average reactor catalyst temperature (to main-tain the design conversion rate) with time is called the catalyst foulingrate. It is an important parameter, used to make an estimate of time whenthe EOR conditions are likely to be reached. The refinery keeps a recordof the reactor average temperature with time, starting from the day thefeed is introduced into the reactor after the new catalyst has been loaded.A graph is drawn between the time on stream vs. the average reactortemperature (Figure 3-7). The data may show scatter, so a straight line isdrawn through the data. From this curve, an estimate of the catalystfouling rate and remaining life of the catalyst can be estimated.

QUENCH BED 2 = MMSCFD

RE

AC

TOR

AV

G.

BE

DBED 2

RE

AC

TOR

AV

G.

BE

D

INO

UT

AVG.

QUENCH BED 3= 18.7 MMSCFD

BED 3

AVG.

RE

AC

TOR

AV

G. B

ED

OU

TIN

QUENCH BED 4 = 18.7 MMSCFD

RE

AC

TOR

AV

G. B

ED

UP

PE

RM

IDLO

WE

R M

ID

DESULFURAZATION BED

BE

D

AVG.

Figure 3-6. Hydrocracker reactor temperature profile.

IN

BED1

OU

T

BE

D1

BE

D 2

BE

D 3

BE

D 4

OU

T

RE

AC

TOR

AV

ER

AG

E T

EM

PE

RA

TUR

ETRENDS IN AVG REACTORTEMPERATURE °F

END OF RUN CATALYST TEMPERATURE 8050F

REACTORAVERAGE TEMPERATURE

PREDICTEDTEMPERATURE

(000) HOURS ONSTREAM

Figure 3-7. Estimating the remaining catalyst life.

EXAMPLE 3-1

The design start of run (SOR) temperature of a hydrocracker reactor is775°F. After the unit is onstream for 12,000 hours, the CAT (weightedaverage bed temperature) is 8000F. Estimate the remaining life of thecatalyst if the design EOR temperature of the catalyst is 8050F:

^ 1 , v (800-775)Catalyst fouling rate = — n o o o

= 0.00208°F/hr

_ . . r f (805 - 800)Remaining life = ^ ^

= 2404 hr or 3.33 months

HYDROGEN PARTIAL PRESSURE

The factors affecting hydrogen partial pressure in the hydrocrackerreactor are

1. Total system pressure.2. Makeup hydrogen purity.3. Recycle gas rate.4. HP gas bleed rate.5. HP separator temperature.

Hydrogen partial pressure in the hydrocracker reactor is a basic oper-ating parameter that provides the driving force for hydrocracking reac-tions. Also, the hydrogen partial pressure has an important effect on thecatalyst fouling rate (Figure 3-8). An increase in the hydrogen partialpressure serve to suppress the catalyst fouling rate. In an operating unit,the hydrogen partial pressure is maximized to operate the unit at thelowest possible temperature. This increases the run length and minimizelight ends production.

FOU

LIN

G R

ATE

T/

Hr

HYDROGEN PARTIAL PRESSURE (psia)

Figure 3-8. Effect of H2 partial pressure on catalyst fouling rate (schematic).

FEED RATE

Increasing the feed rate requires an increase in average catalyst tempera-ture to maintain the required feed rate conversion. The increased feed ratealso causes an increase in the catalyst fouling rate and the hydrogenconsumed chemically and dissolved in the high-pressure separator liquid.The effect of the feed rate on the catalyst fouling rate is shown in Figure 3-9.

FEED CHARACTERIZATION

A heavier feed, as characterized by the ASTM Dl 160 weighted boilingpoint, requires an increase in average catalyst temperature to maintain thedesired level of feed conversion. An increase in the catalyst fouling ratealso occurs. Further, a feed having a higher end point for a given weightedboiling point requires an increased temperature for desired conversionover that required by a lower-end point feed.

LIQUID FEED CONVERSION

The following equation shows how to determine liquid feed conversion:

(reactor liquid feed rate — fractionator bottom rate)Conversion rate = ————

reactor liquid reed rate

RE

LAT

IVE

FO

ULI

NG

RA

TE

ACTUAL

DESIGN

BASIS

CONVERSION 54%

UNITTHROUGHPUT, MBPSD

Figure 3-9. The effect of the feed rate on the catalyst fouling rate (schematic).

RE

LATI

VE

CA

TALY

ST

FOU

LIN

G R

ATE

FOULING RATE AT ACTUAL CONDITIONRELATIVE FOULING RATE =

FOULING RATE AT DESIGN CONDITION

ACTUAL

DESIGN

CONVERSION VOL%

Figure 3-10. The effect of conversion on the catalyst fouling rate (schematic).

The liquid feed conversion depends strongly on temperaturedependent and, as such, has a direct effect on the catalyst fouling rate(Figure 3-10).

GAS BLEED RATE

Bleeding minimizes the buildup of light hydrocarbons in the recyclinggas, which lowers hydrogen partial pressure. Increasing the gas bleed ratelowers the light hydrocarbon concentration in the recycling gas andincreases the hydrogen partial pressure. Decreasing the bleed rate allowslight hydrocarbons in the recycle gas to build up to higher concentrationsand thus lower hydrogen concentration. The bleed rate is typically kept atabout 20% of the hydrogen chemical consumption.

CATALYST SULFIDING AND UNIT STARTUP

The catalyst is sulfided during startup. Before initiating the sulfiding,the distillation and the amine absorber are put into operation.

Sulfiding is the injection of a sulfur-containing chemical such asdimethyl sulfide into the hot circulating gas stream prior to the introduc-tion of liquid feed. The chemical injection is performed with a chemicalinjection pump from a storage drum under an inert gas blanket. Thereactor is first evacuated to 24-26 in. mercury to remove all air, testedfor leaks, and purged with nitrogen a number of times.

The furnace is next purged and fired, and the reactor is heated to4500F. When the reactor is ready for sulfiding, the reactor inlet tempera-ture is increased to 500-5250F. When the catalyst temperature reaches500-525° at the top bed, this temperature is held until the temperature isat least 500° at the inlet and about 450° at reactor outlet.

The feed heater fire is reduced. With recycle compressor operating, thereactor pressure is adjusted to about 300 psi. At 300psig, the hydrogenis introduced slowly, through the makeup compressor and pressureincreased to 700psig at reactor inlet.

Sulfiding chemical or sour gas injection is then added into gas stream at thereactor inlet, at a rate equivalent to 0.5 mol% but no more than 1.0 mol% H2S.

Addition of a sulfiding agent causes two temperature rises in thereactor, first due to reaction of hydrogen with the sulfiding chemical toform H2S, which occurs at the reactor inlet, and second due to reaction ofthe sulfiding chemical with the catalyst, and this moves down through thecatalyst bed. The catalyst temperature is closely monitored during sulfid-ing, and it is not allowed to exceed 6000F.

After the sulfiding temperature rise has passed through the reactor, theH2S concentration of gases out of the reactor begin to rise rapidly. When0.1% H2S is detected in the effluent gas, the recycle gas bleed isstopped and sulfiding continued at a low injection rate until the concen-tration of H2S in the circulating gas is l-2mol%. At the same time, thereactor inlet temperature is increased to 560° and the reactor outlettemperature to at least 535°F.

When no significant reaction is apparent in the reactor, the systempressure is increased by adding makeup hydrogen until a normal operat-ing pressure is reached at the suction of the recycling compressor. Thesulfiding medium is added batchwise to maintain the H2S concentrationof l-2mol% in the recycle gas.

When the reactor system has remained steady at the design pressure,at reactor inlet/outlet temperature of 560/530° and H2S concentration of1-2 mol%, for 2 hours, sulfiding is complete. The reactor is slowly cooledby reducing inlet temperature and adding quench between the catalystbeds until reactor temperature reaches about 425°F. No point in thereactor is allowed to cool below 425°F.

The startup temperature is next approached from the lower side byincreasing the reactor inlet temperature to 450-475° and outlet tempera-ture to 425-450°, the quench temperature controller is set at 450°F. Therecycle gas rate is set at the design rate. Also the HP separator pressure isset at the design value. Condensate and polysulfide injection at the designrate is started.

The reactor is now ready for introduction of the feed. The feed pumpstarts pumping only 20% of the design feed rate to the reactor, shuntingthe rest of the feed back to the feed tank.

Heat is released due to adsorption of hydrocarbon when the feed passesover the catalyst for the first time. This shows up as a temperature wavethat passes down through the catalyst. When the temperature wave due tohydrocarbon adsorption has passed through the reactor and a liquid levelis established in the high pressure separator, the liquid is sent to a low-pressure separator and the distillation section.

As the reactor temperature and feed rate is increased in the stepsthat follow, amine circulation is established through the recycling gasabsorber.

The reactor feed rate is increased by 10-15% of the design at a time,allowing the system to line out for at least 1 hour after each feed rateincrease. The process is repeated until the feed is at 50% of the design rate.

Next, the temperature is increased in about 20°F increments or less atthe inlet to each reactor bed. When increasing reactor temperature, the topbed should be adjusted first then each succeeding lower bed. The systemis allowed to line out for at least 1 hour after each reactor temperatureincrease until all reactor temperatures are steady. Then, the reactortemperature is increased again until the desired conversion is reached.

When the desired operation (conversion and feed rate) has beenreached and the unit is fully lined out, it is monitored so that recyclegas rate is steady and at the design rate. The H2S absorber bypass isslowly closed, forcing all the recycling gas through the absorber. Whileclosing this bypass, close attention is given to knockout drum level,absorber-packed bed AP, and absorber level to avoid compressor shut-down due to a sudden carryover of amine solution.

NORMAL SHUTDOWN

The following procedures are typical when a run is ended to replacespent catalyst or perform general maintenance. The same procedure isfollowed when a run is ended for catalyst regeneration, except that thereactor is not opened.

Care must be taken to avoid furnace and catalyst coking during shut-down and the formation of highly toxic nickel carbonyl when the reactoris cooled, the possibility of fires due to explosive hydrogen oxygenmixtures or exposure of pyrophoric material to air when the reactor isopened, and exposure of personnel to toxic or noxious conditions whenthe catalyst is drained or equipment is entered.

SHUTDOWN PROCEDURE

The following presents a general procedure to be followed for a normalshutdown.

1. Gradually reduce the liquid feed and adjust the catalyst temperaturedownward, so that feed conversion remains at the desired level.When a very low feed rate is reached, stop all liquid feed to thereactor but continue to circulate recycle gas rate at the normalrate. As the liquid feed rate is cut, the reactor feed and effluent rateswill be out of balance for short periods. The feed must be graduallyreduced to prevent temperatures greater than the maximum designtemperature occurring in the effluent/feed heat exchangers.

2. To strip off as much of adsorbed hydrocarbon as possible from thecatalyst, raise the reactor temperature to the normal operating tem-perature after the feed is stopped. Circulate hot hydrogen until nomore liquid hydrocarbon appears in the high-pressure separator. Ifthe shutdown is for only brief maintenance, which does not requirestopping the recycle compressor, continue circulation at the normaloperating temperature and pressure. After the maintenance work iscomplete and unit is ready for feed, lower the reactor temperature andintroduce the feed following the normal startup procedure.

3. If the shutdown is for catalyst regeneration, catalyst replacement,or maintenance that requires stopping the recycle compressoror depressurizing the reactor system, continue circulation at thenormal operating pressure for 2 hours. Reduce the furnace firing

rate and start gradually reducing the reactor temperature to 4000F.Any cooldown design restrictions for the reactor must be adheredto, to avoid thermal shock.

4. Add quench gas as required to evenly cool the reactor.5. To maintain heater duty high enough for good control and speed up

cooling after heater fires are put out, bypass reactor feed gas aroundthe feed/effluent exchanger, as required.

6. While cooling, remove the bulk of hydrocarbon oil from the high-pressure separator by raising the water level and pressure to thelow-pressure separator. Also, increase the pressure of any low-pressure liquid remaining in the H2S absorber to the amine section.Purge and block the liquid hydrocarbon line from the high-pressureseparator. Block in the H2S absorber and drain the amine lines.Drain all vessels and the low points of all lines to remove hydro-carbon inventory.

7. If the shutdown is for catalyst regeneration, hold the reactor at4000F.

8. If the shutdown is for catalyst replacement or maintenance thatrequires opening the reactor, determine the CO content of therecycle gas. If the CO content is below 30ppm CO, continuethe cooling procedure; if more than 30ppm, proceed as follows:

a. If more than 30ppm CO has been detected in the recycle gasand the reactor is to be opened, the CO must be purged beforecooling can continue to eliminate possible formation of metalcarbonyls. Stop recycling gas circulation while the catalyst tem-perature is still above 4000F. Do not cool any portion of the bedbelow 4000F. Before depressurizing the system, check that allthese valves are closed:

• Suction and discharge valves on charge pump and spares.• Block lines on liquid feed lines.• Block valves on the high-pressure separator, liquid product

and water lines.• Recycle and makeup compressor suction and discharge lines.• Chemical injection and water lines.• Makeup hydrogen line.• H2S absorber.

b. Depressure the system. Do not evacuate below atmosphericpressure, as there is the danger of explosion if air is drawn into

the system. For effective purging of the reactor system, thenitrogen gas is injected at the recycle compressor dischargeand follows the normal flow path through the catalyst beds.

c. Pressurize the system with dry nitrogen to 15Opsig. Purge thecompressor with nitrogen.

d. Depressurize the system to 5 psig and purge with nitrogen for 5minutes.

e. Repressurize the system to 150 psig and purge all blocked linesas mentioned previously.

f. Pressurize the system with nitrogen from nitrogen header andrecycle compressor.

9. Relight the furnace fires and maintain the reactor at a temperatureof 4000F for at least 2 hours. Then analyze the recycling gas forCO content. If the CO content is still above 30 ppm, repeat thepreceding procedure.

10. If the CO content is below 30 ppm, the reactor can be cooled below4000F following the cooling rates restrictions for the hydrocrackervessel metallurgy. Stop the furnace fires and continue cooling bycirculating nitrogen until the temperature has fallen to 1200F.

11. Stop nitrogen circulation and ensure that the system is blocked inpreparation to depressurizing. Depressurize and purge the systemagain with nitrogen. Maintain the system under slight nitrogenpressure, so that no oxygen is admitted.

CATALYST REGENERATION

The activity of the catalyst declines with time on stream. To recover thelost activity, the catalyst is regenerated in place at infrequent intervals.

In the regeneration process, the coke and sulfur deposits on the catalystare burned off in a controlled manner with a dilute stream of oxygen. Theburning or oxidation consists of three burns at successively higher tem-peratures. This is followed by a controlled reduction of some of thecompounds formed during the oxidation step, in a dilute hydrogen stream.The catalyst is next stabilized by sulfiding. During the combustion phase,sulfur oxides are liberated and some metal sulfates are formed as thesulfided catalyst is converted to the oxide form. During the reductionphase, the residual metal sulfates are reduced to the metal sulfides withliberation of sulfur dioxide.

During both the combustion and the reduction phases, a dilute causticquenching solution is mixed with reactor effluent to cool the gasesthrough the H2O-SO3 dew point and to neutralize the SO2/SO3 present.Cooling the gases in this manner prevents corrosion of heat exchangersand other downstream equipment. Also, sulfur oxides must be neutralizedto prevent damage to the catalyst.

During regeneration, maximum use is made of equipment used in thenormal operation. The recycle compressor is used to recirculate the inertgas. Compressed diluted makeup air is mixed with recycled inert gasflowing to the reactor inlet to provide oxygen to sustain a burn wave inthe top catalyst beds. The reactor inert gas is preheated in the feedeffluent exchanger then combined with diluted makeup air and heatedin the feed furnace before flowing to the reactor inlet.

The reactor/effluent gases are cooled in the feed effluent exchangerand mixed with recirculating dilute caustic quench solution beforebeing cooled in the effluent/air cooler. The cooled mixture then flowsto the high-pressure separator to separate out the vapor and liquid phases.A portion of the inert gas from the separator is bled to the atmosphere andthe reminder recirculated. The dilute caustic quench solution then flowsto the low-pressure separator. This allows the use of a lean DEA pump forquencher solution circulation. Some of the dilute caustic quench solutionis bled to limit the concentration of dissolved solids. Fresh causticand process water are added to maintain the quench solution at a properpH.

The regeneration is carried out at the pressure that allows the max-imum flow rate as limited by the maximum power of the recycle com-pressor or the discharge temperature of the makeup compressor(1150psig). The makeup air is diluted with the recycling inert gas toavoid forming a combustible mixture in the lubricated makeup compres-sor. With hydrogen present the maximum O2 concentration is 4mol%maximum at a compressor outlet temperature of 265°F.

After the regeneration burn has started and the recycling gas stream isfree of hydrogen, the O2 concentration through the makeup compressor israised to 7%.

The concentration of O2 in the first burn is kept at 0.5%. The initialburn is started at an inlet temperature of 6500F or less. This produces atemperature rise of 10O0F. The burn is concluded when the last bedtemperature starts to drop off and oxygen appears in the recycle gas.

After the first burn is completed, the reactor inlet temperature isincreased to 7500F and air is introduced to obtain 0.5 vol% oxygen at

the reactor inlet. The oxygen may not be totally consumed as it passesthrough the reactor during this burn and the oxygen content of recyclinggas at the reactor outlet may remain at fairly constant level, even thoughoxygen is being consumed in the reactor. This burn normally producesonly a small temperature rise and sometimes no temperature rise isapparent.

For the third or final burn, the temperature at reactor inlet is raised to8500F ± 25°F and the oxygen content of the inlet gas is slowly increasedto 2% by volume. This condition is maintained at least for 6 hours or untilthe oxygen consumption is essentially nil, as indicated by a drop in inletoxygen content of less than 1% for 2 hours with makeup air stopped. Nopoint in the reactor should be allowed to exceed 875°F. The reactortemperature should never be allowed to exceed 9000F. Throughout theregeneration, a circulating dilute caustic solution is injected into reactoreffluent gases.

REDUCTION PHASE

After the hydrocarbon, coke, and sulfur have been burned from thecatalyst, a reduction step is necessary before introducing the sulfiding andfeed. During oxidation, a portion of sulfur on the catalyst is convertedinto sulfates. These sulfates may be reduced when exposed to high-pressure hydrogen, and the reduction can occur at temperatures belowthose required for sulfiding. A large heat release accompanies thesereactions, which can result in loss of catalyst activity or structural damageto the catalyst and reactor.

The oxidized catalyst is reduced in a very dilute hydrogen atmosphereso that the heat of reduction is released gradually under controlledconditions.

The reactor is pressurized with nitrogen while heating to 6500F.Hydrogen is then introduced to obtain a concentration of l-2mol% atthe reactor inlet. This causes a controlled reduction reaction to movethrough the reactor, as indicated by a temperature rise of 25^K)0F perpercent hydrogen.

If, after 1 or 2 hours, the reduction is proceeding smoothly, slowlyincrease concentration of hydrogen to 2-3 mol% at the reactor inlet. If thetemperature rise exceeds 1000F or the reactor temperature at the reductionwave exceeds 7500F, stop increasing the hydrogen content. Do notexceed 3% hydrogen at the reactor inlet or allow the temperature toexceed 7500F at any point in the reactor during this procedure.

The reduction reaction is controlled by limiting the quantity of hydro-gen available. The catalyst temperature is held below 7500F during thereduction step to avoid damage to the catalyst.

When hydrogen breaks through the bottom of the reactor, reduce orstop hydrogen addition. Hold the reactor temperature at 650° and 2-3%hydrogen at the reactor inlet for at least 8 hours and until hydrogenconsumption has dropped to a low rate. When the wave has passedthrough the reactor and no more hydrogen is consumed, the reactor iscooled to about 4000F.

Some sulfur oxides are generated by the reduction reactions and con-sequently a dilute quench is circulated as during the oxidation phase toneutralize the quenched effluent gases. However, since the quantity ofsulfur oxides is relatively small compared to that liberated during thecarbon/sulfur burn, an ammonia quench may be used instead of a causticone. This arrangement uses equipment used during normal operationrather than the quench circulation, piping, and mixing header. It isimportant to note that a catalytic reformer hydrogen is unsuitable forreduction and the first phase of subsequent sulfiding. Accordingly, theprocedure requires manufactured or electrolytic hydrogen.

MILD HYDROCRACKING

Mild hydrocracking, as the name suggests, operates at much lowerpressure and much milder other operating conditions than the normalhydrocracking process. The objective of the process is basically to desul-furize the VGO to make it suitable for FCCU feed. Other impurities, likenitrogen, are also removed and about 30% of the feed is converted intosaleable diesel. Compared to normal hydrocracker units, mild hydro-crackers require much less initial investment. The operating conditionsand yield from a mild hydrocracker unit are shown in Tables 3-5 and 3-6,respectively.

RESIDUUM HYDROCRACKING

Resid hydrocracking is designed to convert straight-run residual stocksand atmospheric or vacuum resids into distillates by reacting them withhydrogen in the presence of a catalyst. The process operates under severeoperating conditions of high temperature and pressure, comparable to

Table 3-6Mild Hydrocracker Yields

VGO FEED 1.0000HYDROGEN 0.0061

TOTAL FEED 1.0061

GASES 0.0191CRACKED NAPHTHA 0.0134KEROSENE 0.0110DIESEL 0.3117HEAVY DIESEL 0.6509

TOTAL PRODUCTS 1.0061

NOTES: ALL FIGURES ON A WAV BASIS.FEED SG = 0.9169, 2.7% SULFURPRODUCT.DIESEL: 0.07% S, 43 DI.HEAVY DIESEL: 0.12% S.

Table 3-7Resid Hydrocracking Catalyst Characteristics

PROPERTY UNITS

CATALYST COMPOSITION NI-MO ON ALUMINA BASESHAPE EXTRUDATES 0.963 mm x 3.93 mmBULK DENSITY lb/ft3 45CATALYST DENSITY, lb/ft3 105

WITH PORESCATALYST DENSITY, lb/ft3 186

EXCLUDING PORES

distillate hydrocracker units. A part of the resid is converted into dis-tillates. Also, the resid is partially desulfurized and demetallized. Thedistillates produced are separated in a distillation column into naphtha,kerosene, and light diesel. The heavy products from the reaction sectionand fractionation tower are separated in a vacuum distillation section intoheavy diesel, heavy vacuum gas oil, and vacuum resid (see Tables 3-7and 3-8). The major reactions that occur are shown in Figure 3-11.

RESID HYDROCRACKER REACTOR

The resid hydrocracking reactions are conducted in an ebullated orfluidized bed reactor to overcome the problems associated with a fixed

Table 3-8Resid Hydrocracking Spent

Catalyst Composition

COMPONENT UNITS WT%

NICKEL Wt% 2.1MOLYBDENUM Wt% 3.73COBALT Wt% 0.006CARBON Wt% 12.18SULFUR Wt% 11.9VANADIUM Wt% 7.1

bed (see Figure 3-12). A liquid phase passes upward through a bed ofcatalyst at a velocity sufficient to maintain the catalyst particles in contin-uous random motion. This liquid velocity is achieved by circulating a liquidrecycling stream by means of an ebullating pump external to the reactor.

An ebullating bed system offers the following advantages over theconventional fixed-bed system:

(1) CRACKING AND HYDROGENATION

CATALYST

(2) HYDRODESULFURISATION ANDDENITRIFICATION

CATALYST

Figure 3-11. Resid hydrocracking reactions.

CATALYST

CATALYSTADDITION

MAX LIQ. LEVEL

LEVEL DETECTORS EXPANDED LEVEL

SETTLED CATALYSTLEVEL

GAS/LIQUIDPRODUCTTOSEPARATORS

CATALYSTWITHDRAWAL

MAKEUP H2

AND FEED OIL

DISTRIBUTORGRID PLATE

RECYCLINGOIL

EBBULATIONPUMP

Figure 3-12. Resid hydrocracker reactor.

1. Isothermal reactor conditions. The mixed conditions of this reactorprovide excellent temperature control of the highly exothermicreactions without the need for any quench system. Undesirabletemperature sensitive reactions are controlled.

2. Constant pressure drop. Since the catalyst is in a state of constantrandom motion, there is no tendency for pressure drop to build up asa result of foreign material accumulation.

3. Catalyst addition and withdrawal. The catalyst can be added orwithdrawn from an ebullating bed on either a continuously or inter-mittent basis. This feature permits operation at an equilibrium activ-ity level, thereby avoiding change in yield and product quality withtime encountered in fixed-bed reactors due to aging of the catalyst.

Within the reactor, the feed enters the lower head of the reactorthrough a sparger to provide adequate distribution of the reactant stream.The ebullating stream is distributed through the lower portion of thereactor by an individual sparger. These spargers effect a primary distribu-tion of the feed stream across the reactor cross-sectional area. The feedthen passes through a specially designed distributor plate, which furtherensures uniform distribution as the vapor and liquid flow upward throughthe catalyst bed.

The oil and hydrogen dissolved in liquid phase under relatively highhydrogen partial pressure react with each other when brought in intimatecontact with active catalyst above the distribution plate. The primaryreactions taking place are hydrocracking, hydrogenation, hydrodesulfur-ization, and denitrification. In addition, the organometallic compounds inthe feed are broken down under high temperature and high hydrogenpartial pressure and are, in part, adsorbed on the catalyst, the remainderpassing through the catalyst bed, ultimately ending up in fuel oil. Themetal buildup on the catalyst would result in complete deactivation of thecatalyst. Therefore, the activity level of the catalyst is maintained byaddition of fresh catalyst and withdrawal of spent catalyst in a pro-grammed manner.

The temperature and the catalyst activity level control the conversionlevel. The other variables such as hydrogen partial pressure, circulatinggas rate, reactor space velocity, and ebullating rate are unchanged in anoperating unit.

The reactor average temperature is varied by the amount of preheatingperformed on the total oil and gas streams that pass through separate firedheaters. Normally, the oil heater outlet is maintained at a moderatetemperature to minimize skin cracking of the oil, and the adjustment ofthe reactor temperature is done primarily by increasing the preheating ofthe hydrogen feed gas to the maximum temperature possible within theheater design limitation.

The ebullating oil flow is controlled by varying the speed of theebullating pump. Within the reactor, the ebullating liquid is drawn intoa conical collecting pan, located several feet above the catalyst bedinterface to ensure catalyst-free liquid circulation down the internal standpipe and into external ebullating pump. The ebullating liquid is distrib-uted through the reactor bottom by its individual sparger.

The height of the fluidized bed of the catalyst in the reactor is related tothe gas flow rate, the liquid flow rate, and the physical properties of thefluid, which in turn are affected by the operating temperature and pres-sure, conversion, size, density, and shape of the catalyst particles. Allthese parameters are maintained within the constraints imposed by thesecorrelations.

RESID HYDROCRACKER UNITS

The feed to resid hydrocracker is crude unit vacuum bottoms. Freshfeed is mixed with a heavy vacuum gas oil diluent. The combined feedenters surge drum V-33, is pumped to feed heater H-Ol then mixed withhigh-pressure recycled hydrogen gas, is preheated in H-02 in mixers, thenfed to the reactor V-Ol at about 7500F and 2450psig (see Figure 3-13).

Ebullation pump P-02 maintains circulation to keep catalyst particlesin suspension. The catalyst bed level is monitored by a radioactive source.Fresh catalyst is added to the reactor and spent catalyst is withdrawn fromthe reactor periodically to maintain the catalyst level.

The reactor effluent flows as a vapor/liquid mixture into vapor/liquidseparator V-02. The vapor from separator drum is a hydrogen-rich streamcontaining equilibrium quantities of hydrocarbon reaction products. Theliquid is composed essentially of heavier hydrocarbon products fromresid hydrocracking reactions and unconverted feed. The hot flash liquidcontains an appreciable amount of dissolved hydrogen and light ends. Theflash vapor leaving V-02 is cooled with light distillate in exchanger E-02.The cooled vapor effluent is then fed to primary distillate knockout drumV-03 for removal of condensed liquid. Vapor from V-03 is further cooledto about 1400F in air cooler E-03 and then enters distillate separator drumV-04. On leaving separator V-04, the hydrogen-rich gas is split into twostreams. The larger stream is compressed to approximately 2700 psig ina recycle gas compressor and returned to the reactor V-Ol as recycledgas. The smaller purge gas stream is withdrawn as purge to maintain thepurity of the recycled gas.

LP FLASH GAS

27 PSIG

2510 PSIG

140'F

RECYCLEING GAS

2530 PSIG

FLASH

TREAT GASHEATER

FEED SURGE

MAKEUP

KEROSENE

E-07

1500F

HEATER

P-08

Figure 3-13. Resid hydrocracking (reactor section).

Condensed liquid from V-03 joins with reactor effluent liquid fromV-02, then the combined stream is flashed to remove hydrogen in threesuccessive stages.

The first stage let down occurs at 1050 psig in V-07. The resultinghydrogen-rich vapor is cooled in exchanger E-05 then cooled with cool-ing water to approximately 1100F in E-06. This cooled stream is flashedagain in flash gas drum V-16. The vapor from V-16 is let down toa pressure of about 550 psig and joins the high-pressure flash gas fromV-05. The liquid from V-16 is further let down to a pressure of 550 psig inhigh-pressure flash drum V-05.

Ammonical condensate is injected upstream of E-03 and E-06 todissolve ammonium bisulfide deposits. The ammonical water fromV-04 and V-05 are returned to the ammonical water treatment plant tostrip away H2S and NH3.

The second stage of let down of the combined liquid stream to about550 psig occurs in high-pressure, high-temperature flash drum V-14. Vaporfrom this flash drum is cooled with light distillate in exchanger E-12 andcooling water in E-18. It is then fed to flash drum V-05, in addition tohigh-pressure condensate from V-04 and V-16. High-pressure flash gasfrom V-05 and gas let down from V-16 are sent to an amine wash.

The final let down of the combined reactor liquid stream occurs at75 psig in steam stripper V-09. The resulting vapor is cooled through lightdistillate exchanger E-19 to 3000F and then through a water cooler to theambient temperature before it is fed to low-pressure flash drum V-06 atabout 30psig, along with condensate from V-05. The low-pressure flashgas from V-06 is sent to the amine wash unit.

The light distillate from V-06 is pumped by P-03 through exchangersE-19 and E-12, E-05, and E-02, preheated to approximately 6000F, andfed to the fractionator V-10.

The hot flashed reactor liquid from V-14 flows into stripper V-09,where it is steam stripped to remove middle distillates. The strippedoverhead vapors are fed to fractionator tower V-10 along with preheatedlight distillate from V-06.

The fractionator tower feed consists of light and middle distillates fromV-06 and stripped gases from V-09. The fractionator produces twosidestreams and overhead vapor and liquid streams. A light diesel cut isfed to diesel side stripper V-13 for removal of light ends by steamstripping. The raw light diesel stream is pumped by P-06 and cooled inair cooler E-07/8 before sending it to storage.

A kerosene cut is withdrawn to stripper V-12 for removal of lightcomponents by steam stripping. The raw kerosene stream is sent by P-07,cooled in air cooler E-07, and sent to storage.

A pumparound sidestream is circulated near the top of the fractionatorto provide internal reflux to upper section with P-08 and air cooler E-IO.

A lower pumparound is taken at light diesel drawoff by pump P-05,passing through air cooler E-07 to the fractionator to provide internalreflux in the lower section of the tower.

The fractionator overhead vapor is cooled and partially condensed inair cooler E-Il before going to naphtha accumulator drum V-08. Vaporfrom this drum is sent to an amine wash. The unstabilized liquid naphthafrom V-08 is sent to storage by pump P-07.

The separated water withdrawn from V-08 is sent to the oily watercondensate system.

PRODUCT FRACTIONATION

The fractionator bottom product is pumped by P-16 to a vacuum feedsurge drum, then the charge is heated to a temperature of about 7800Fbefore it enters the vacuum tower (see Figure 3-14). The flash zone is at7000F and 26 inch vacuum maintained by steam ejectors.

The feed is fractionated into heavy diesel, heavy vacuum gas oil, andvacuum bottoms.

The heavy diesel is withdrawn as side cut, cooled in air cooler E-52,and sent to storage. The heavy vacuum gas oil (HVGO) side cut is cooledby raising 150 psig steam then in air cooler E-54. A part of HVGO is usedas diluent in the resid hydrocracker feed, and the rest is used in fuel oilblending.

The vacuum bottoms from the unit are pumped by P-53 to off-sitestorage tanks and used in fuel oil blending.

The typical operating conditions of resid hydrocracker units are shownin Table 3-9, and yields and product qualities are shown in Tables 3-10to 3-12.

The liquid hourly space velocity is calculated based on the flowvolume of the liquid feed at the reactor operating conditions divided bythe volume of the catalyst up to the fluidized bed height. Only the freshfeed flow is considered, since the net recycle flow is zero. Also, thevolume of the makeup and recycled hydrogen is neglected.

SOUR WATER

BOILER FEEDWATER

HEAVY DIESEL

HVGO

VACUUM RESID

STEAMDRUMV-52

E-51

E-52

E-54E-53

E-55

P-52

P-53

BFW CIRCULATION

P-51

VACUUMTOWER

V-51

155°F

625°F

455°F

QUENCHING

700TPRESSURE26 INCH Hq

6030F

P-54

7500F

P-55

Figure 3-14. Residuum hydrocracking (distillation section). BFW = boiler feedwater.

HEATERH-51

FLAMEARRESTOR

STEAM

SEAL POTV-50

EJECTORS

EJ-01.02

STRIPPERBOTTOMS

Table 3-9Resid Hydrocracker Operating Conditions

OPERATING VARIABLE UNIT

REACTOR AVERAGE TEMP 0F 827REACTOR SYSTEM PRESSURE psig 2400HYDROGEN PARTIAL PRESSURE psig 1895MAKEUP HYDROGEN scf/bbl 1400RECYCLING GAS FLOW scf/bbl 4300BLEED GAS FLOW scf/bbl 700HYDROGEN CONSUMPTION scf/bbl 660LIQUID HOURLY SPACE VELOCITY (hr"1) 0.906HP FLASH GAS PRESSURE psig 500LP FLASH GAS PRESSURE psig 43FEED/H2 MIX TEMPERATURE 0F 748PLENUM CHAMBER TEMPERATURE 0F 811CATALYST ADDITION RATE lb/bbl 0.106

EXAMPLE 3-2

Calculate the LHSV for a resid hydrocracker reactor with the followingoperating conditions:

Resid feed rate = 24,150 bpsd

Diluent in feed = 3000 bpsd

Total reactor feed = 27,150bpsd, 600F

Reactor cross section = 133 ft2

Reactor temperature = 8500F

Fluidized bed height = 47 ft

Specific gravity of oil at 600F = 0.900

Specific gravity at 8500F = 0.553

Makeup H2 flow = 33.7mmscf

Recycled H2 flow = 69.5mmscf

Neglecting the volume of the makeup and recycled hydrogen gas, thesuperficial liquid linear velocity in the reactor is due to total temperature

Table 3-10Resid Hydrocracker Feed and Product Qualities

VACUUMBOTTOMVGO

HEAVYDIESEL

LIGHTDIESELKEROSENENAPHTHAFEEDUNITSPROPERTY

6.61.0245

7.320.8

800880925

301170

41203.59363

18.90.9408

615

740760815850905945970

20002

24.80.9065

440540575600710765785800820

18751.6

30.90.8713

20

365468520548592622688720744

12950.97

40.90.8208

22

318360372380422452490504522

7290.43

64.50.7219

30

102144158184244278310322344

1550.19

7.51.0186.8517.2

725870930975

321989

25754.82775

Wt%Wt%

ppmw

ppmwWt%Cst, 2100F

API GRAVITYDENSITYASPHALTENESCON CARBONBROMINE NUMBERDISTILLATION, ASTM

IBP5%10%20%50%70%90%95%EP

METALSNiNaVNITROGENSULFUR

VISCOSITY

Table 3-11Resid Hydrocracker Yields

STREAM YIELD WT FRACTION

FEED 0.8840DILUENT 0.1160HYDROGEN 0.0360

TOTAL 1.0360

PRODUCTSGASES 0.0626H2S 0.0200NAPHTHA 0.0600KEROSENE 0.0913LIGHT DIESEL 0.0655HEAVY DIESEL 0.0797HEAVY GAS OIL 0.1833RESIDUE 0.4647LOSSES 0.0089

TOTAL 1.0360

corrected volume of fresh oil plus that of recycled oil. Since the netrecycle flow is zero, this can be neglected.

Flow rate in the reactor = 27,150 x 0.90 x 5.6146/(24 x 0.553)

= 10,337 ft3/hr

Reactor velocity = 10,337/(133 x 3600)

= 0.0215 ft/sec

Catalyst volume up to fluidization level = 47 x 133

= 6251 ft3

Liquid hourly space velocity = 6251/10,337

= 0.6047

From the data in Table 3-10, we see that virtually all products from theunit require some further treatment before blending. Naphtha, kerosene,and diesels have high bromine numbers and high nitrogen content, requir-ing further hydrofinishing treatment to make them suitable for blending.Similarly, the heavy diesel cut requires further hydrotreating to reduce itsnitrogen content and make it suitable for FCCU or hydrocracker feed.

Table 3-12Utilities Consumption (per Ton Feed)

UTILITY UNITS CONSUMPTION

ELECTRICITY kWh 5.70FUEL mmBtu 1.06STEAM mmBtu 0.20COOLED WATER MIG* 1.56CATALYST LBS 0.71

*MIG = 1000 IMPERIAL GALLONS.